Multiple riser reactor with centralized catalyst return

ABSTRACT

The present invention is directed to a hydrocarbon conversion apparatus and process. The apparatus comprises the following: a plurality of riser reactors, each having a first end into which a catalyst is fed, a second end through which the catalyst can exit, and optionally a center axis extending therebetween. The apparatus also includes a separation zone having a plurality of inlets, each inlet not being oriented along the center axes of the riser reactors, the separation zone being provided to separate the catalyst from products of a reaction conducted in the hydrocarbon conversion apparatus. A plurality of deviating members are also provided, each deviating member being in fluid communication between the second end of a respective riser reactor and a respective inlet of the separation zone. The apparatus also includes a catalyst retention zone provided to contain catalyst, which is fed to the riser reactors. A catalyst return is in fluid communication between the separation zone and the catalyst retention zone.

[0001] This application claims priority to U.S. Ser. No. 60/419,408,filed Oct. 18, 2002, which is fully incorporated herein by reference.The present invention relates to a reactor useful in hydrocarbonconversion processes and particularly in oxygenate to olefin conversionreactions.

FIELD OF THE INVENTION BACKGROUND OF THE INVENTION

[0002] When converting a feedstock containing a hydrocarbon to a productin an industrial reactor, it is desirable to maximize the production ofa desired product or products, and to control, typically to minimize,the production of by-products. One type of reactor useful for conductinghydrocarbon conversion reactions is a fluidized bed reactor, whereinsolid catalyst particles are suspended in a fluidized state duringcontact with the feedstock and other vapor materials. These types ofreactors usually have a cylindrical reactor geometry. One method forreducing the production of by-products in a fluidized bed reactorinvolves operating in a hydrodynamic flow regime such that thesuperficial gas velocity obtains a velocity high enough to cause the netflow of catalyst in the reactor to flow in the same direction as theflow of the feedstock and other vapors, i.e., the feedstock and othervapors essentially carry the catalyst particles along with them. Theseflow regimes are known to those skilled in the art as the fast-fluidizedbed and riser regimes, more generally as the transport regime, and arepreferred in reaction systems in which a more plug flow reactor type isdesired.

[0003] In general, for a given reactor cross sectional area (which in acylindrical reactor geometry is proportional to the diameter, and moregenerally to a characteristic width), the catalyst concentration in afluidized bed reactor decreases with increasing gas superficialvelocity. Higher gas superficial velocities generally require tallerreactor heights to allow a given amount of feedstock to contact arequired amount of catalyst. These higher gas superficial velocitiesnecessitate a higher aspect ratio (the ratio of a reactor height to itsdiameter or characteristic width) of the reactor. Further, in many casesit is desired to make a fluidized reactor with a very largecross-sectional area to enable very large throughputs of feedstock in asingle reactor facility. However, increasing fluid bed diameter,particularly in the transport regime, also necessitates increasedreactor height. This increased height is required because a certainminimum reactor height, in terms of a minimum aspect ratio, is requiredto achieve a fully developed flow pattern, which approximates plug flowreactor behavior. At the exit and, particularly, at the entrance of atransport regime fluidized bed reactor, unsteady state momentum effectsdominate hydrodynamic behavior (e.g., the energy required for thefeedstock vapors to pick up and accelerate the solid catalyst againstthe force of gravity) in a manner not conducive to obtaining approximateplug flow behavior. Not until these momentum effects are dampened out byprogressing along the reactor height will a well behaved, approximatelyplug flow fluid/solid flow pattern emerge. Finally, should the use oflower activity catalysts be desired in the transport regime, aspectratios must also increase to provide desired higher feedstockconversion.

[0004] Unfortunately, high aspect ratio transport fluid bed reactors aredifficult and expensive to construct and maintain. They are expensivebecause they must have at the top a very large, heavy separation vessel,often filled with heavy equipment, to capture and manage the flowingcatalyst and reactor product. As the reactor increases in height (aspectratio), more expensive support structures may be required. In certainareas of the world where inclement, especially windy weather occursroutinely, even more structural support is required, and certain aspectratios are not economic. Multiple, complete and independent reactorsystems with independent separation vessels are required. With thesemultiple, complete and independent reactor systems come attendantmultiplication of costs.

[0005] Thus, a need exists in the art for a reactor that can provide thedesired aspect ratio without necessitating an unwieldy height, forcing awidth in which the desired, fully developed flow regime may never beobtained, or without resulting to multiple, independent reactor systems.

SUMMARY OF THE INVENTION

[0006] The present invention provides a solution to the currentlyexisting need in the art by providing a shorter hydrocarbon conversionapparatus while maintaining a high aspect ratio in the riser reactors.The hydrocarbon conversion apparatus includes a plurality of riserreactors, each of which is in fluid communication with a respectivedeviating member, which deviates the flow of product, unreacted feed,and catalyst from the top of the riser reactor to a plurality of sideinlets on a separation zone. Preferably, the inlets are formed in theupper section of the separation zone. By deviating the flow of materialfrom the top of the riser reactors through inlets oriented in the uppersection of the separation zone, the overall height of the separationzone, and thus the height of the overall hydrocarbon conversionapparatus, may be advantageously reduced.

[0007] Additionally, by orienting the riser reactors adjacent thevertically extending sides of the separation zone, rather than orientingthe riser reactors beneath the separation zone, the present inventionprovides the additional advantage of increasing the number of possibleriser reactors in a given hydrocarbon conversion apparatus. Byincreasing the number of riser reactors in a given hydrocarbonconversion apparatus, the quantity of product formed per unit time isincreased.

[0008] In one embodiment of the present invention, the hydrocarbonconversion apparatus comprises a plurality of riser reactors, eachhaving a first end into which a catalyst may be fed, a second endthrough which the catalyst can exit, and a center axis or centroid lineextending therebetween. A separation zone having a plurality of inlets,each inlet preferably not being oriented along the center axes of theriser reactors, is provided to separate the catalyst from products of areaction conducted in the hydrocarbon conversion apparatus. A pluralityof deviating members is also provided, each deviating member being influid communication between the second end of a respective riser reactorand a respective inlet of the separation zone. The apparatus alsoincludes a catalyst retention zone provided to contain catalyst, whichis then fed to the riser reactors, and a catalyst return in fluidcommunication between the separation zone and the catalyst retentionzone.

[0009] The invention is also directed to a hydrocarbon conversionapparatus including a plurality of riser reactors, each having a firstend into which a catalyst is fed and a second end through which thecatalyst can exit the riser reactor. A separation zone is provided toseparate the catalyst from products of a reaction conducted in thehydrocarbon conversion apparatus, the separation zone including aplurality of inlets. The apparatus also includes a plurality ofdeviating members, each deviating member being oriented to deviate aflow of material from the second end of a respective riser reactortoward a respective inlet. At least one catalyst return is in fluidcommunication with the separation zone and the first ends of the riserreactors, the catalyst return being provided to transfer the catalystfrom the separation zone to the first ends of the riser reactors.

[0010] The invention is also directed to a hydrocarbon conversionprocess including: (a) contacting a fluidizable catalyst with afluidizing agent to fluidize the fluidizable catalyst; (b) feeding thecatalyst and a feed to a plurality of riser reactors, the plurality ofriser reactors being part of a single hydrocarbon conversion apparatus;(c) contacting the feed with the catalyst in the plurality of riserreactors under conditions effective to convert the feed to a product;(d) directing the product and the catalyst through a plurality ofdeviating members each deviating member being positioned to deviate aflow of the product and the catalyst from an outlet of a respectiveriser reactor to a separation zone; (e) separating the catalyst from theproduct in the separation zone, the separation zone being in fluidcommunication with the plurality of deviating members; (f) returning thecatalyst from the separation zone to the plurality of riser reactors;and (g) repeating steps (a) to (f).

[0011] In another embodiment, the invention is to a hydrocarbonconversion apparatus including a plurality of riser reactors, eachhaving a first end for receiving catalyst and a second end through whichthe catalyst and a product can exit. A plurality of deviating membersare provided, each deviating member being associated with a respectiveriser reactor. A separation zone having at least one side is provided toseparate the catalyst from the product, wherein the separation zoneincludes a plurality of inlets, each inlet being associated with arespective deviating member, and wherein the inlets are oriented on theside of the separation zone. The apparatus also includes a catalystreturn coupled to a plurality of arms, the catalyst return and armsbeing in fluid communication between the separation zone and the firstends of the plurality of riser reactors.

BRIEF DESCRIPTION OF THE DRAWINGS

[0012] This invention will be better understood by reference to theDetailed Description of the Invention when taken together with theattached drawings, wherein:

[0013]FIG. 1 presents a partial cross sectional view of a hydrocarbonconversion apparatus of the present invention;

[0014]FIG. 2 presents a partial cross sectional view of anotherembodiment of the hydrocarbon conversion apparatus of the presentinvention;

[0015] FIGS. 3A-3C present partial cross sectional top views of threeembodiments of hydrocarbon conversion apparatuses according to thepresent invention;

[0016]FIG. 4 presents a cut-away side view of an elbow deviating memberaccording to one embodiment of the present invention;

[0017]FIG. 5 presents a cut-away partial cross sectional side view of ariser reactor closed-coupled with a cyclone separator;

[0018]FIG. 6 presents a partial cross sectional view of a deviatingmember providing a deviation angle of less than 90 degrees; and

[0019]FIG. 7 presents a partial cross sectional view of a “J-Bend”catalyst return system according to another embodiment of the presentinvention.

DETAILED DESCRIPTION OF THE INVENTION

[0020] The present invention provides a hydrocarbon conversion apparatusincluding a plurality of riser reactors, each of which is in fluidcommunication with a respective deviating member. Each deviating memberdeviates the flow of product, unreacted feed, and catalyst from the topof a respective riser reactor to one of a plurality of inlets on theseparation zone. Preferably the inlets are formed on the side of theupper section of the separation zone. By deviating the flow of materialfrom the top of the riser reactors through inlets oriented in the uppersection of the separation zone, the overall height of the separationzone, and thus the overall height of the hydrocarbon conversionapparatus, may be advantageously reduced. The present invention providesthe additional advantage of allowing the size of the separation vesselto be reduced because the top of the riser reactors are not required toextend inside the separation vessel.

[0021]FIG. 1 presents a partial cross sectional view of a hydrocarbonconversion apparatus (HCA), generally designated 110, in accordance withone embodiment of the present invention. The HCA 110 comprises a shell160, separation zone 124, a plurality of riser reactors 114, a feeddistributor 154, and a catalyst return 130. With continuing reference toFIG. 1, the shell 160 defines the separation zone 124 in which a productof the hydrocarbon conversion reaction is separated from the catalyst,which catalyzes the hydrocarbon conversion reaction. Shell 160 andseparation zone 124 include a first end 162 and a second end 128. Theseparation zone 124 preferably includes one or more separation devices126, which are used to separate the products from the catalyst. Theseparation devices 126 may be cyclonic separators, filters, screens,impingement devices, plates, cones or any other device, which wouldseparate the catalyst from the product of the conversion reaction. Theseparation devices 126 shown in FIG. 1 are cyclonic separators. In otherembodiments, not illustrated, the separation devices are positionedexternally to the separation zone 124, outside of the shell defining theseparation zone 124, or a combination of externally and internallypositioned separation devices.

[0022] Riser reactors 114 extend adjacent to shell 160 and theseparation zone 124. Each riser reactor 114 includes a first end 116into which the catalyst and feed are fed to conduct the hydrocarbonconversion reaction. If the riser reactor 114 is a cylindrical form, asshown in FIG. 1, a center axis 158 extends between the first end 116 andthe second end 118. If the riser reactor lacks a central axis, e.g., isnot cylindrical, conical, etc., a centroid line perpendicular to across-sectional plane of the riser reactor extends between the first andsecond ends. Preferably, the separation zone 124 is a cylindrical formhaving a center axis 166, which is preferably parallel to the centeraxis 158 of riser reactors 114. Each riser reactor 114 further includesa second end 118 through which the catalyst, products and unreactedfeed, if any, exit the riser reactor 114. The first end 116 of eachriser reactor 114 terminates in a mouth 164 through which the catalystis fed into the riser reactor 114. The number of riser reactors 114employed in the HCA 110 varies depending on the hydrocarbon conversionprocess to be conducted in the apparatus 110. The apparatus 110 cancontain two, three, four, five, six or even more than six riser reactors114.

[0023] The geometry of the riser reactors 114 may vary widely. In theembodiment in FIGS. 1 and 2, the geometry is comprised of a cylindricalmember. Optionally, the geometry of the riser reactors is comprised of asingle, right frustum of a cone. Other suitable geometric shapesinclude, but are not limited to, triangular prisms and frusta ofpyramids, rectangular and square wedges and frusta of pyramids, andpentagonal, hexagonal, septagonal and octagonal prismatoids and frustathereof, general and right. Further non-limiting examples includevarious polyhedrons, such as a tetrahedron, an octohedron, adodecahedron or an icosahedron, and conical spheres and sphericalsectors, and torus and barrels in their circular, elliptical orparabolic forms, and frusta thereof, general and right. Multipleoccurrences of any of these geometric shapes defining the riser reactorand/or the associated separation zone are also within the scope of anapparatus of the present invention.

[0024] In accordance with the present invention, at least one riserreactor 114, preferably a plurality of riser reactors 114 are orientedexternally to the shell 160. If shell 160 is substantially cylindrical,riser reactors 114 are positioned beyond the maximum diameter defined byshell 160. Riser reactors 114 preferably are not positioned directlyunderneath separation zone 124. In one embodiment, the riser reactors114 are positioned adjacent to separation zone 124. In this embodiment,the second ends 118 of each riser reactor 114 enter separation zone 124above the second end 128 of separation zone 124.

[0025] Second end 118 of each riser reactor 114 preferably is in fluidcommunication with a deviating member 120. The deviating member 120preferably is a curved pipe, elbow, or other conduit device adapted toreceive catalyst, products and unreacted feed, if any, that exits riserreactor 114 through second end 118. The deviating member 120 deviatesthe flow of the catalyst, products and unreacted feed from the secondend of the riser reactor 114 toward inlet 122 on the side wall(s) ofshell 160. Preferably, the inlet 122 is not oriented along the centeraxis 158 (or centroid line of a riser reactor lacking a center axis) ofriser reactor 114. Preferably, the deviating member provides a deviatingangle of greater than 0 degrees. Preferably, at least one of thedeviating members has a deviation angle of at least 10, at least 20, atleast 30 or at least 45 degrees, more preferably at least 60, at least80 or at least 90 degrees. The deviation angle is defined herein as theangle that the flow of catalyst, products and unreacted feed, if any, isdeviated by deviating member 120 from the second end 118 of riserreactor 114 to inlet 122. In FIG. 1, the deviation angle is illustratedby deviation angle θ, which is defined by center axis 158 of a riserreactor (or corresponding centroid line of a riser reactor lacking acenter axis) and the entry flow path, shown by arrow 170. In FIG. 1, thedeviation angle θ is approximately 90 degrees.

[0026]FIG. 6 illustrates another embodiment, wherein the deviatingmember assembly, generally designated by arrow 600, provides a deviationangle of less than 90 degrees. As shown in FIG. 6, a riser reactor 602having a center axis 612 (or centroid line if the riser reactor lacks acenter axis) is in fluid communication with a deviating member 604,which deviates the flow of catalyst, product and unreacted feed, if any,to inlet 608 defined by shell 610. Shell 610 also defines separationzone 614, as discussed above. However, unlike the previously-describedembodiments, the deviation angle is less than 90 degrees. In FIG. 6, thedeviation angle λ is defined by center axis 612 of a riser reactor 602(or corresponding centroid line of a riser reactor lacking a centeraxis) and the entry flow path, shown by arrow 606. In FIG. 6, thedeviation angle λ is approximately 45 degrees. Erosion is advantageouslyreduced by providing a deviating member, which defines a deviation angleof less than 90 degrees. In terms of ranges, in one embodiment, thedeviation angle is from about 10 to about 90 degrees, more preferablyfrom about 30 to about 90 degrees and most preferably from about 45 toabout 90 degrees. In other embodiments, the deviation angle is greaterthan 90, greater than 100, or greater than 120 degrees.

[0027] One embodiment of the present invention, not shown, includes asecond plurality of riser reactors oriented underneath the separationzone. Each of the second plurality of riser reactors includes first andsecond ends and a second center axis extending therebetween. Theseparation zone includes a plurality of second inlets, each second inletbeing oriented along a respective second center axis. Each of the secondplurality of riser reactors optionally extends inside the separationzone. In this embodiment, the number of riser reactors for a givenseparation zone may be greatly increased. Accordingly, the amount ofproduct formed in a single hydrocarbon conversion apparatus can also besignificantly increased.

[0028] Reverting to FIG. 1, the size of the riser reactors 114 dependson parameters such as superficial gas velocity, solids hydrodynamics,pressure, and production capacity of the desired hydrocarbon conversionprocess. In the present invention, each riser reactor 114 desirably hasa height from 10 meters to 70 meters and a width (or diameter) of onemeter to three meters. All of the riser reactors 114 have a similarheight from their first ends 116 through their second ends 118.Desirably, the heights of the riser reactors 114 vary by no more than20% from one riser reactor 114 to another riser reactor 114. Moredesirably, the heights vary by no more than 10% and, most desirably, theheights vary by no more than 1%.

[0029] In one embodiment of the present invention, each of the riserreactors 114 has a similar cross sectional area along its entire height.Desirably, each of the riser reactors has a cross sectional area of nogreater than 12 m². More desirably, each of the riser reactors has across sectional area of no greater than 7 m². Most desirably, each ofthe riser reactors has a cross sectional area of no greater than 3.5 m².Desirably, the cross sectional areas of the riser reactors vary by nomore than 20% from one riser reactor to another riser reactor. Moredesirably, the cross sectional areas of the riser reactors vary by nomore than 10% and, most desirably, the cross sectional areas of theriser reactors vary by no more than 1%. If one or more riser reactorshave both a largest and a smallest cross-sectional area at differentpoints along the height of riser reactors, desirably the largestcross-sectional areas of the riser reactors vary by no more than 20%from one riser reactor to another riser reactor, and the smallestcross-sectional areas of the riser reactors vary by no more than 20%from one riser reactor to another riser reactor. More desirably, thelargest cross sectional area of one riser reactor varies by no more than10% from the largest cross sectional area of another riser reactor andthe smallest cross sectional area varies by no more than 10% from thesmallest cross sectional area of another riser reactor. Most desirably,the largest cross sectional area of one riser reactor varies by no morethan 1% from the largest cross sectional area of another riser reactorand the smallest cross sectional area varies by no more than 1% from thesmallest cross sectional area of another riser reactor. Preferably eachof the riser reactors has a width (or diameter) of from 1 to 3 meters.

[0030] Desirably, the cross sectional area of each riser reactor variesby no more than 50% along its entire height. More desirably, the crosssectional area of each riser reactor varies by no more than 30% alongits entire height and, most desirably, the cross sectional area of eachriser reactor varies by no more than 10% along its entire height.

[0031] In one embodiment, to provide a feed to the riser reactors 114,at least one feed distributor 154 is positioned near the first ends 116of the riser reactors 114. The feed distributor 154 receives a feedstream from a feed providing line 150 and directs the feed to one ormore of the riser reactors 114. In another embodiment, not shown, morethan one feed distributor 154 is employed adjacent the first ends 116 ofthe riser reactors 114 to provide feed in various states, e.g., one feeddistributor may provide feed in a vapor form while a second feeddistributor may provide feed in a liquid form. Each feed distributor 154includes a body from which a plurality of necks 152 extend. Each riserreactor 114 has at least one associated neck 152. Each neck 152terminates in a head 112. Each head 112 of each neck 152 is positionedadjacent to the first end 116 of each riser reactor 114. Desirably, eachhead 112 extends upwardly into each riser reactor 114. More desirably,each head 112 is positioned at or above the mouth 164 at the first end116 of each riser reactor 114. Feed distributor 154 may include anoptional flow control device, not shown, positioned on feed distributor154 to control the amount of feed to each neck 152 or a flow controldevice may be positioned on each neck 152. The flow control device canalso be employed to measure flow as well as control it. Further, anozzle, not shown, may be positioned on each head 112 to further controlthe distribution of the feed to each riser reactor 114. Additionally,each head 112 may be fitted with a screening, plunger or other device,not shown, to prevent back flow of catalyst into any of necks 152 offeed distributor 154.

[0032] At least one catalyst return 130 provides fluid communicationbetween the separation zone 124 of shell 160 and the riser reactors 114.The apparatus 110 may include one, two, three, four, five, six or morecatalyst returns 130, although only a single catalyst return 130 isillustrated in FIGS. 1 and 2. If a plurality of catalyst returns aredesired, each of the catalyst returns preferably is adapted to delivercatalyst from the separation zone directly to each of the respectiveriser reactors. Typically, although not necessarily, a single catalystreturn 130, which is centrally oriented with respect to the riserreactors 114, is used. In this embodiment, the single catalyst return isin fluid communication with a plurality of arms 136. As shown in FIG. 1,a single catalyst return 130 is positioned centrally in relation to theriser reactors 114. The catalyst return 130 has a first end 140 and asecond end 142. The first end 140 of the catalyst return 130 opens intothe second end 128 of shell 160 and the second end 142 of catalystreturn 130 opens to a series of arms 136 adapted to deliver catalyst tothe first ends 116 of riser reactors 114.

[0033] The arms 136 extend from the catalyst return 130 to each of theriser reactors 114 and provide fluid communication between the catalystreturn 130 and the riser reactors 114. Each arm includes a first end 168adjacent the catalyst retention zone 134 and catalyst return 130, and asecond end 148 adjacent the riser reactor 114. Catalyst flows througheach arm 136 from the first end 168 to the second end 148. The number ofarms 136 preferably will correspond to the number of riser reactors 114with each riser reactor having at least one corresponding arm 136. Flowof catalyst through the catalyst return 130 optionally is controlledthrough the use of flow control device(s) 144 positioned on the catalystreturn 130 and/or on each arm 136. The flow control devices may be anytype of flow control devices currently in use in the art to controlcatalyst flow through catalyst transfer lines. If employed, the flowcontrol device 144 is desirably a ball valve, a plug valve or a slidevalve.

[0034] In the embodiment shown in FIG. 1, the second end 142 of thecatalyst return 130 and the arms 136 define a catalyst retention zone134. The arms 136 open to the catalyst retention zone 134. The catalystretention zone 134 is provided to retain catalyst, which is used tocatalyze the hydrocarbon conversion reaction, which is conducted in theapparatus 110. As one of skill in the art will appreciate, the boundarybetween the catalyst retention zone 134 and the catalyst return 130 isfluid and depends, at least in part, on the level of catalyst containedin the catalyst retention zone 134 and the arms 136.

[0035] At least one fluidizing agent distributor 132 is positionedbeneath the catalyst retention zone 134. The fluidizing agentdistributor 132 includes a conduit into which a fluidizing agent is fedto fluidize a fluidizable catalyst in the catalyst retention zone 134and the catalyst return 130. Additional fluidizing agent distributors132, as shown in FIG. 1, may also be positioned on the catalyst return130 and/or on one or more of the arms 136 to further fluidize catalystcontained therein. Optionally, the catalyst retention zone 134 includesa disperser, not shown, positioned in the catalyst retention zone andprovided to disperse the fluidizing agent in the catalyst retention zoneto facilitated fluidization of the catalyst therein. For example, thedisperser could be a device selected from the group consisting of agrid, a screen and a perforated plate. Preferably, at least onedisperser extends in a plane perpendicular to the center axis 166 of theseparation zone 124 (or centroid line if the separation zone lacks acenter axis) and is oriented above one or more fluidizing agentdistributors 132.

[0036]FIG. 7 illustrates a “J-Bend” catalyst return system, generallydesignated by numeral 700, in accordance with another embodiment of thepresent invention. As shown in FIG. 7, catalyst return 712 is shown influid communication with an optional catalyst retention zone 710, whichis similar to the catalyst retention zone 134 illustrated in FIG. 1.Catalyst return 712 also is in fluid communication with a plurality ofstandpipes or arms 708, each of which is provided to transfer catalyst,preferably in a fluidized manner, from the catalyst return 712 and/orcatalyst retention zone 710 to the first end of a respective riserreactor 706. More specifically, each of the arms 708 is adapted todeliver catalyst to a U-shaped member 714, which deviates the flow ofcatalyst from the second end of a respective arm 708 toward the firstend of a respective riser reactor 706. Optionally, each of the arms 708includes a flow control device, not shown, as discussed above withreference to FIG. 1. Preferably, one or more fluidizing agentdistributors 702 are provided adjacent to one or more of the catalystretention zone 710, the arms 708 and the U-shaped member 714, in orderto provide a fluidizing agent to maintain the catalyst in a fluidizedstate. In this embodiment, the oxygenate feedstock preferably isintroduced into the sides of the riser reactor 706 through sidefeedstock introduction nozzles 704. In this manner, the flow of catalystfrom the arms to the riser reactors is not restricted by the presence ofone or more necks 152, illustrated in FIGS. 1 and 2.

[0037] Reverting to FIG. 1, the hydrocarbon conversion apparatus 110 mayalso include an outlet and outlet line 146 through which the catalyst isremoved from the apparatus 110, e.g., for catalyst regeneration. Theoutlet line 146 is shown as being positioned on the second end 128 ofthe shell 160 but may be positioned at any position on the apparatus110. Thus, the hydrocarbon conversion apparatus 110 of the presentinvention optionally includes an associated catalyst regenerationapparatus, not shown. The catalyst regeneration apparatus is in fluidcommunication with the hydrocarbon conversion apparatus 110. Thecatalyst regeneration apparatus includes a catalyst regenerator, whichis in fluid communication with the hydrocarbon conversion apparatus 110,and an optional catalyst stripper, not shown, which is in fluidcommunication with the catalyst regenerator and which is in fluidcommunication with the hydrocarbon conversion apparatus 110. A firstline, shown in part by outlet line 146, provides fluid communicationbetween the catalyst stripper and the outlet on shell 162. A secondline, not shown, provides fluid communication between the catalyststripper and the catalyst regenerator. A third line, shown in part asinlet line 156, provides fluid communication between the catalystregenerator and the inlet on shell 160. A flow control device, notshown, optionally is positioned on first line to control the flow ofcatalyst between the shell 160 and the catalyst stripper. Additionallyor alternatively, a flow control device, not shown, is positioned on thesecond line to control the flow of catalyst between the catalyststripper and the catalyst regenerator. Additionally or alternatively, aflow control device, not shown, is positioned on the third line tocontrol the flow of catalyst between the catalyst regenerator and theshell 160. The flow control devices may be any types of flow controldevices currently in use in the art to control catalyst flow throughcatalyst transfer lines. Useful flow control devices include ballvalves, plug valves and slide valves. The catalyst stripper optionallyis separate from or integrally formed with the catalyst regenerator. Thethird line can return catalyst to any portion of the HCA 110. Forexample, in various embodiments, catalyst is returned to the separationzone 124, as shown in FIG. 1, to the catalyst return 130, the catalystretention zone 134, the arms 136, directly to riser reactor 114, or anycombinations thereof.

[0038] The apparatus 110 shown in FIG. 1 functions in the followingmanner. The apparatus 110 is filled with an appropriate amount ofcatalyst suitable to carry out the desired hydrocarbon conversionreaction. The catalyst should be of a type that is fluidizable. At leasta portion of the catalyst is retained in the catalyst return 130 and thecatalyst retention zone 134. The catalyst is fluidized in the catalystreturn 130 and the catalyst retention zone 134 by means of a fluidizingagent, which is provided to the hydrocarbon conversion apparatus 110through the conduits of the fluidizing agent distributors 132. Usefulfluidizing agents include, but are not limited to, inert gasses,nitrogen, steam, carbon dioxide, hydrocarbons and air. The choice offluidizing agent depends upon the type of conversion reaction beingconducted in the HCA 110. Desirably the fluidizing agent is unreactive,e.g., inert, in the reaction being conducted in the HCA 110. In otherwords, it is desirable that the fluidizing agent does not play a part inthe hydrocarbon conversion process being conducted in the HCA 110 otherthan to fluidize the fluidizable catalyst.

[0039] Once the catalyst has reached an acceptable fluidized state, afeed is fed into the HCA 110 through feed distributor 154. The feedenters the body of feed distributor 154, passes through the necks 152 offeed distributor 154 and exits through the heads 112 of feed distributor154. The feed is distributed to each of the riser reactors 114 throughtheir first ends 116. Desirably, the feed is provided in substantiallyequal streams to each riser reactor 114. By “substantially equal” it ismeant that the flow of feed provided to each riser reactor 114 throughthe feed distributor 154 varies by no more than 25% by volume rate, andvaries no more than 25% by mass percent, for each component in the feed,from one riser reactor to another riser reactor. More desirably, theflow of feed provided to each riser reactor 114 through the feeddistributor 154 varies by no more than 10% by volume rate, and varies nomore than 10% by mass percent for each component in the feed, from oneriser reactor to another riser reactor. Most desirably, feed provided toeach riser reactor 114 through the feed distributor 154 varies by nomore than 1% by volume rate, and varies no more than 1% by mass percentfor each component in the feed, from one riser reactor to another riserreactor.

[0040] A pressure differential created by the velocity of the feedentering the first ends 116 of the riser reactors 114 and the pressureof the height of fluidizable catalyst in the catalyst return(s) 130 andthe catalyst retention zone 134 causes catalyst to be aspirated into thefirst ends 116 of the riser reactors 114. The catalyst is transportedthrough the riser reactors 114 under well known principles in which thekinetic energy of one fluid, in this case the feed, is used to moveanother fluid, in this case the fluidized catalyst. The catalyst andfeed travel from the first ends 116 to the second ends 118 of the riserreactors 114. As the catalyst and feed travel through the riser reactors114, the hydrocarbon conversion reaction occurs and a conversion productis produced. The flow of catalyst to the riser reactors 114 arecontrolled by the flow control devices 144.

[0041] By designing the hydrocarbon conversion apparatus 110 with thesefeatures, each individual riser reactor 114 operates in a substantiallyidentical manner. With this invention, it is desirable to maintain boththe reactant feed rates and the catalyst feed rates at the same rates toeach of the riser reactors 114. In this way, the conversion of the feedand selectivity to the desired products will be substantially identicaland can run at optimum operational conditions.

[0042] The conversion product(s), unreacted feed, if any, and thecatalyst exit the riser reactors 114 through their second ends 118 andenter the deviating member 120. The deviating member preferably is acurved pipe, elbow, or other conduit device adapted to receive catalyst,products and unreacted feed, if any, that exits riser reactor 114through second end 118. The deviating member 120 deviates the flow ofthe catalyst, products and unreacted feed from the second end of theriser reactor 114 toward inlet 122 on the side wall(s) of separationzone 124. The conversion product(s), unreacted feed, if any, andcatalyst then enters the separation zone 124 of shell 160. In theseparation zone 124, the conversion product and unreacted feed, if any,are separated from the catalyst by a separation devices 126, such ascyclonic separators, filters, screens, impingement devices, plates,cones, other devices that would separate the catalyst from the productof the conversion reaction, and combinations thereof. Desirably, asshown in FIG. 1, the conversion product and unreacted feed, if any, areseparated by a series of cyclonic separators. Once the catalyst has beenseparated from the conversion product and the unreacted feed, if any,the conversion products and unreacted feed, if any, are removed from theshell 160 through the product exit line 138 for further processing suchas separation and purification.

[0043] Product exit conduit or conduits 176 from separation device(s)126 is openly joined to a plenum shell 172. Plenum volume 174 is formedwithin the boundaries of plenum shell 172 as joined to the section ofthe shell 160 defining the top of separation zone 124. The plenum shell172 and plenum volume 174 are provided to collect reaction product andpossibly unreacted feedstock exiting separation devices 126 via productexit conduit or conduits 176, and direct that material to product exitline 138. The product exit line 138 is openly joined to separation zone124 in the vicinity of plenum volume 174, and is provided to conveyreaction product and possibly unreacted feedstock away from theapparatus. Such plenum designs are particularly useful in embodimentswhen a plurality of separation devices are utilized, for example asdisclosed in FIGS. 1 and 2. As shown, the product exit conduits 176 fromthe separation devices 126 are openly joined to the plenum shell 172,and a single, secondary product exit line 138 may be used to carryproduct away from the hydrocarbon conversion apparatus.

[0044] The catalyst, after being separated from the products andunreacted feed, moves from the shell 160 to catalyst return 130 and thecatalyst retention zone 134. The catalyst exits shell 160 through thefirst end 140 of the catalyst return 130 and moves through the catalystreturn 130 to the second end 142 of the catalyst returns 130 from whichthe catalyst moves to the optional catalyst retention zone 134. Ifdesired, the flow of catalyst through the catalyst returns 130 iscontrolled by the flow control devices 144. If the flow control devices144 are employed, a height of fluidizable catalyst is maintained aboveeach flow control device 144 in the catalyst return 130 to allow properfunction of the flow control device 144.

[0045] As discussed above, if necessary or desired, at least a portionof the catalyst is circulated to the catalyst regeneration apparatus.Catalyst to be regenerated is removed from the shell 160 through theoutlet and outlet line 146 and transported, if desired, to the catalyststripper. Optionally, the flow of catalyst between the hydrocarbonconversion apparatus 110 and the catalyst stripper is controlled by aflow control device, not shown. In the catalyst stripper, the catalystis stripped of most of readily removable organic materials (organics).Stripping procedures and conditions for individual hydrocarbonconversion processes are within the skill of a person of skill in theart. The stripped catalyst is transferred from the catalyst stripper tothe catalyst regenerator through a second line, not shown. The flow ofcatalyst through the second line optionally is controlled by one or moreflow control devices. In the catalyst regenerator, carbonaceous depositsformed on the catalyst during a hydrocarbon conversion reaction are atleast partially removed from the catalyst. The regenerated catalyst isthen transferred to the shell 160 of the hydrocarbon conversionapparatus 110 through a third line. The flow of catalyst through thethird line optionally is controlled by one or more flow control device.A transport gas is typically provided to the third line to facilitatetransfer of the catalyst from the catalyst regenerator to thehydrocarbon conversion apparatus 110. The catalyst is returned to theshell 160 through the inlet 156.

[0046]FIG. 2 illustrates a hydrocarbon conversion apparatus similar tothe one shown in FIG. 1 having a quiescent zone 204 to facilitatecatalyst removal. In this embodiment, the catalyst return 130 includes awall extension 202, which extends upwardly into the second end 128 ofshell 160, and a funnel portion 208. The wall extension 202 and thefunnel portion 208 define a quiescent zone 204 in which a portion of thecatalyst is retained prior to being removed from the shell 160 viaoutlet 146. As catalyst accumulates in the quiescent zone 204 up to theedge of wall extension 202, excess catalyst will spill out of thequiescent zone 204 and into catalyst return 130.

[0047]FIG. 2 also illustrates an embodiment of the present inventionwherein the first ends 168 of arms 136 act as the catalyst retentionzone. In this embodiment, the plurality of arms 136 optionally meet atan apex 206 rather than meeting in a separate and distinct catalystretention zone 134, illustrated in FIG. 1. In the embodiment shown inFIG. 2, catalyst flow to the riser reactors 114 is controlled by flowcontrol devices 144, discussed above. An apex fluidizing agentdistributor 210 optionally is provided adjacent the apex 206 tofacilitate catalyst flow into arms 136, as shown in FIG. 2. It will beappreciated that the HCA of FIG. 2 operates in a similar manner to theHCA described above with regard to FIG. 1.

[0048] Representative embodiments of possible configurations of riserreactors and separation zones in accordance with the present inventionare shown by top views in FIGS. 3A-3C. Although each of FIGS. 3A-3Cillustrates an HCA 110 having four riser reactors 114, it is, to beunderstood that these configurations could be modified to provide 2, 3,4, 5, 6 or even more than 6 riser reactors similarly oriented aboutseparation zone 124.

[0049]FIG. 3A shows a possible configuration for the riser reactors 114for the HCA 110 shown in FIG. 1. As shown in FIG. 3A, the riser reactors114 are oriented externally and adjacent to separation zone 124 definedby shell 160. Each riser reactor 114 is coupled to a deviating member120, which is adapted to deviate the flow of catalyst, product andunreacted feed, if any, from the top of riser reactors 114 toward aninlet 122 on the side wall of the shell 160 defining the separation zone124. Each of the deviating members 120 deviates the flow of catalyst,product and unreacted feed, if any, from the top of the riser reactors114 toward the separation zone 124, preferably toward the center axis166 (or centroid line), as shown in FIG. 1, of the separation zone 124.Optionally, each deviating member is a tubular member arcuously curvedabout a single plane, which preferably passes through the center axis166 (or centroid line) of the separation zone 124 and the center axis158 (or centroid line) of the riser reactor 114, as shown in FIG. 1. Theflow of material from the top of the riser reactors 114 is illustratedby arrow 300, which preferably is oriented perpendicular to an imaginarytangent line on the cylindrical outer surface of the shell 160 definingseparation zone 124 adjacent the center of the inlet 122. If the shelland separation zone are not oriented in a cylinder, the flow of materialis preferably oriented substantially perpendicular to the surface of theshell 160 adjacent inlet 122. In this embodiment, the catalyst, productand any unreacted feed from riser reactors on opposite sides of theseparation zone 124 may collide with each other, or more likely withseparation devices 126, shown in FIG. 1, thereby facilitating mixing ofthese components and providing increased reaction time for thehydrocarbon conversion reaction, which may continue to occur inside theseparation zone 124.

[0050]FIG. 3B shows another embodiment of the present invention whereinseparation of the products of the hydrocarbon conversion reaction fromthe catalyst is facilitated by the formation of a cyclone within theseparation zone. In this embodiment, the cyclone is formed by the flowof material entering the separation zone at relatively high superficialgas velocities from the riser reactors via arcuous deviating members. Asshown in FIG. 3B, the riser reactors 310 are oriented externally andadjacent to separation zone 124 defined by shell 160. Each riser reactor310 is coupled to an arcuous deviating member 306, which is adapted todeviate the flow of catalyst, product and unreacted feed, if any, fromthe top of riser reactors 310 toward the separation zone 124.Optionally, each deviating member 306 is a tubular member arcuouslycurved about a plurality of planes. The flow of material from the top ofthe riser reactors 310 is illustrated by arrow 302, which preferably isoriented at an oblique angle, e.g., not perpendicular, to an imaginarytangent line on the cylindrical outer surface of the shell 160 definingseparation zone 124 at the inlet 314. Accordingly, each inlet 314 isoffset from its respective riser reactor 310. Each of the arcuousdeviating members 306 deviates the flow of catalyst, product andunreacted feed, if any, illustrated by arrow 302, from the top of theriser reactors 310 toward a peripheral region of the separation zone 124via inlet 314. “Peripheral region” means a region of the separation zone124 other than the region defined by the center axis 166, shown in FIG.1, or central region thereof. Preferably, each arcuous deviating member306 directs the flow from its respective riser reactor 310 to asimilarly situated respective peripheral region of the separation zone.By directing the catalyst, product and unreacted feed to similarlysituated respective peripheral regions of the reactor, a cyclone isformed within the separation zone 124. The formation of a cyclonefacilitates catalyst separation in a manner similar to conventionalcyclone separation units, illustrated as separation devices 126 in FIGS.1 and 2. By forming a cyclone within the separation zone, the number ofseparation devices 126 within the separation zone 124 may beadvantageously reduced while maintaining similar or better catalystseparation over conventional HCA's.

[0051]FIG. 3C shows another embodiment of the present invention whereinseparation of the products of the hydrocarbon conversion reaction fromthe catalyst is facilitated, as with FIG. 3B, by the formation of acyclone within the separation unit. The cyclone is formed by the flow ofmaterial entering the separation zone 124 at relatively high superficialgas velocities from the riser reactors via deviating members 308. Asshown in FIG. 3C, the riser reactors 312 are oriented externally andadjacent to separation zone 124 defined by shell 160. Additionally, theriser reactors 312 are offset to a side of the reactor relative to theembodiments illustrated in FIGS. 3A and 3B. Each riser reactor 312 iscoupled to a deviating member 308, which is adapted to deviate the flowof catalyst, product and unreacted feed, if any, from the top of riserreactors 312 toward the separation zone 124. Optionally, each deviatingmember 308 is a tubular member curved about a single plane, which ispreferably parallel to a second plane passing through the axis of theseparation zone 124. The flow of material from the top of the riserreactors 312 is illustrated by arrow 304, which preferably is orientedat an oblique angle, e.g., not perpendicular, to an imaginary tangentline on the cylindrical outer surface of the shell 160 definingseparation zone 124 at inlet 316. Each of the deviating members 308deviates the flow of catalyst, product and unreacted feed, if any,illustrated by arrow 304, from the top of the riser reactors 312 towarda peripheral region of the separation zone 124. Preferably, eachdeviating member directs the flow from its respective riser reactor 312to a similarly situated respective peripheral region of the separationzone 124. By directing the catalyst, product and unreacted feed tosimilarly situated respective peripheral regions of the reactor, acyclone is formed within the separation zone 124.

[0052] Although FIGS. 1-3 illustrate HCA's having deviating membersformed of tubular members, one or more of the deviating members can alsobe formed as an elbow, as illustrated in FIG. 4. In this embodiment,each riser reactor 402 is in fluid communication with an elbow deviatingmember, generally designated 400, which includes a contacting plate 408,a catalyst contact region 412 and a delivery member 404. After catalyst,product and unreacted feedstock, if any, has traveled up the riserreactor 402 it enters catalyst contact region 412. Some catalystcollides at relatively high gas superficial velocities with the innersurface 410 of contacting plate 408. Ideally, the gas superficialvelocity is high enough so that an amount of catalyst is temporarilyretained in catalyst contact region 412. As more catalyst flows up riserreactor 402, the catalyst contacts the retained catalyst in the catalystcontact region 412, and is deviated thereby into delivery member 404.Delivery member 404 is adapted to receive catalyst, product andunreacted feedstock, if any, from the catalyst contact region 412 anddirect the catalyst, product and unreacted feedstock toward the inlet414 formed in shell 160. The delivery member 404 optionally is formed ofa pipe or tubular member, or any other suitable shape for deliveringcatalyst, product and unreacted feed, if any, from the catalyst contactregion 412 to inlet 414. By contacting retained catalyst particles inthe catalyst contact region 412, rather than contacting a wall of thedeviating member, erosion in the deviating member 400 is advantageouslyminimized. The angle formed by the center axis of riser reactor 402 andthe axis of delivery member 404 is preferably about 90 degrees, althoughother angles are possible.

[0053] As shown in FIGS. 1 and 2, the deviating member 120 directs thecatalyst, product and unreacted feed, if any from the riser reactor 114to an inlet 122 in the side wall of the shell 160 defining separationzone 124. That is, the catalyst, product and unreacted feed areintroduced directly into the separation zone 124. A first portion of thecatalyst, product and unreacted feed stock enters an inlet in one ormore of the separation devices 126 wherein the catalyst and product areseparated, e.g., by centrifugation. A second portion of the catalystfalls to catalyst return 130 without being processed by the separationdevices 126.

[0054]FIG. 5 illustrates another embodiment of the present invention,generally designated by numeral 500, wherein all of the material fromthe deviating member is sent directly to one or more separation devices506 thereby ensuring efficient separation of catalyst from product ofthe hydrocarbon conversion reaction. That is, one or more riser reactors504 are closely-coupled with a separation device 508 within, or without,separation zone 124. As shown in FIG. 5, riser reactor 504 is in fluidcommunication with a deviating member 502, which directs catalyst,product and unreacted feed, if any, through shell 160, which definesseparation zone 124. However, unlike the previously-describedembodiments, the deviating member 502 optionally extends inside theseparation zone 124 and directs the catalyst, product and unreactedfeed, if any, to inlet 508 on separation device 506. Although FIG. 5illustrates the deviating member 502 in fluid communication with twocyclonic separators in series, any of the types of catalyst separatorsdescribed above may be implemented in this embodiment. In the separationdevice 506, catalyst is efficiently separated from the product of thehydrocarbon conversion reaction. Optionally, each respective riserreactor and deviating member combination is equipped with its ownseparation device. Alternatively, a plurality of deviating members froma plurality of associated riser reactors can direct catalyst, productand unreacted feedstock, if any, to a single separation device.Optionally, because the riser reactor of this embodiment isclosely-coupled with one or more separation devices, the separationdevice may be external to separation zone 124. If the separation deviceis external to the separation zone, then a separation zone may beadvantageously reduced in size. Additionally, more cyclones may beincorporated in the present invention if they are oriented externally tothe separation zone, without having to enlarge the size of theseparation zone.

[0055] While the riser reactors and catalyst returns are shown in thevarious Figures as having a circular cross section, the riser reactorsand catalyst returns may have any cross section, which would facilitateoperation of the hydrocarbon conversion apparatus. Other useful crosssections for the riser reactors and the catalyst returns includeelliptical cross sections, polygonal cross sections and cross sectionsof sections of ellipses and polygons. Desirable cross-sections for theriser reactors and catalyst returns include circles and regular polygonswith sides of equal lengths. By “regular”, it is meant that the shape ofthe cross-section has no line segments with vertices, inside theboundaries of the shape, having angles greater than 180°. The mostdesirable cross-sections are circles, and triangles, squares, andhexagons with sides of equal length. The means of determiningcross-sectional areas for any cross-section shape is based on longestablished geometric principles well known to those skilled in the art.Similarly, desirable cross-sections for the separation zone includecircles and regular polygons with sides of equal lengths. The mostdesirable cross-sections are circles, and triangles, squares, andhexagons with sides of equal length.

[0056] While the position of the riser reactors relative to theseparation zone are shown in the figures as equidistant and symmetrical,alternate configurations are within the scope of the present invention.For example, the riser reactors may be positioned on one side of theseparation zone in a hemispherical layout. As another example, when theseparation zone has a circular or approximately circular cross-section,the riser reactors may be positioned in a line along the diameter theseparation zone. One skilled in the art will appreciate that a widevariety of configurations of the risers relative to the separation zonemay be utilized in the present invention.

[0057] The hydrocarbon conversion apparatus of the present invention isuseful to conduct most any hydrocarbon conversion process in which afluidized catalyst is employed. Typical reactions include, for example,olefin interconversion reactions, oxygenate to olefin conversionreactions, oxygenate to gasoline conversion reactions, malaeic anhydrideformulation, vapor phase methanol synthesis, phthalic anhydrideformulation, Fischer Tropsch reactions, and acrylonitrile formulation.

[0058] The hydrocarbon conversion apparatus of the present invention isparticularly suited for conducting an oxygenate to olefin conversionreaction. In an oxygenate to olefin conversion reaction, an oxygenate isconverted to an olefin by contacting an oxygenate feed with a catalystunder conditions sufficient to convert the oxygenate to an olefin.

[0059] The process for converting oxygenates to light olefins employs afeed including an oxygenate. As used herein, the term “oxygenate” isdefined to include, but is not necessarily limited to, hydrocarbonscontaining oxygen such as the following: aliphatic alcohols, ethers,carbonyl compounds (aldehydes, ketones, carboxylic acids, carbonates,and the like), and mixtures thereof. The aliphatic moiety desirablyshould contain in the range of from about 1-10 carbon atoms and moredesirably in the range of from about 1-4 carbon atoms. Representativeoxygenates include, but are not necessarily limited to, lower molecularweight straight chain or branched aliphatic alcohols, and theirunsaturated counterparts. Examples of suitable oxygenates include, butare not necessarily limited to the following: methanol; ethanol;n-propanol; isopropanol; C₄-C₁₀ alcohols; methyl ethyl ether; dimethylether; diethyl ether; di-isopropyl ether; methyl formate; formaldehyde;di-methyl carbonate; methyl ethyl carbonate; acetone; and mixturesthereof. Desirably, the oxygenate used in the conversion reaction isselected from the group consisting of methanol, dimethyl ether andmixtures thereof. More desirably the oxygenate is methanol. The totalcharge of feed to the riser reactors may contain additional components,such as diluents.

[0060] One or more diluents may be fed to the riser reactors with theoxygenates, such that the total feed mixture comprises diluent in arange of from about 1 mol % and about 99 mol %. Diluents that may beemployed in the process include, but are not necessarily limited to,helium, argon, nitrogen, carbon monoxide, carbon dioxide, hydrogen,water, paraffins, other hydrocarbons (such as methane), aromaticcompounds, and mixtures thereof. Desired diluents include, but are notnecessarily limited to, water and nitrogen.

[0061] A portion of the feed may be provided to the reactor in liquidform. When a portion of the feed is provided in a liquid form, theliquid portion of the feed may be either oxygenate, diluent or a mixtureof both. The liquid portion of the feed may be directly injected intothe individual riser reactors, or entrained or otherwise carried intothe riser reactors with the vapor portion of the feed or a suitablecarrier gas/diluent. By providing a portion of the feed (oxygenateand/or diluent) in the liquid phase, the temperature in the riserreactors is controllable. The exothermic heat of reaction of oxygenateconversion is partially absorbed by the endothermic heat of vaporizationof the liquid portion of the feed. Controlling the proportion of liquidfeed to vapor feed fed to the reactor is one possible process forcontrolling the temperature in the reactor and in particular in theriser reactors.

[0062] The amount of feed provided in a liquid form, whether fedseparately or jointly with the vapor feed, is from about 0.1 wt. % toabout 85 wt. % of the total oxygenate content plus diluent in the feed.More desirably, the range is from about 1 wt. % to about 75 wt. % of thetotal oxygenate plus diluent feed, and most desirably the range is fromabout 5 wt. % to about 65 wt. %. The liquid and vapor portions of thefeed may be the same composition, or may contain varying proportions ofthe same or different oxygenates and same or different diluents. Oneparticularly effective liquid diluent is water, due to its relativelyhigh heat of vaporization, which allows for a high impact on the reactortemperature differential with a relatively small rate. Other usefuldiluents are described above. Proper selection of the temperature andpressure of any appropriate oxygenate and/or diluent being fed to thereactor will ensure at least a portion is in the liquid phase as itenters the reactor and/or comes into contact with the catalyst or avapor portion of the feed and/or diluent.

[0063] Optionally, the liquid fraction of the feed may be split intoportions and introduced to riser reactors a multiplicity of locationsalong the length of the riser reactors. This may be done with either theoxygenate feed, the diluent or both. Typically, this is done with thediluent portion of the feed. Another option is to provide a nozzle,which introduces the total liquid fraction of the feed to the riserreactors in a manner such that the nozzle forms liquid droplets of anappropriate size distribution, which, when entrained with the gas andsolids introduced to the riser reactors, vaporize gradually along thelength of the riser reactors. Either of these arrangements or acombination thereof may be used to better control the temperaturedifferential in the riser reactors. The means of introducing amultiplicity of liquid feed points in a reactor or designing a liquidfeed nozzle to control droplet size distribution is well known in theart and is not discussed here.

[0064] The catalyst suitable for catalyzing an oxygenate-to-olefinconversion reaction includes a molecular sieve and mixtures of molecularsieves. Molecular sieves may be zeolitic (zeolites) or non-zeolitic(non-zeolites). Useful catalysts may also be formed from mixtures ofzeolitic and non-zeolitic molecular sieves. Desirably, the catalystincludes a non-zeolitic molecular sieve. Desired molecular sieves foruse with an oxygenate to olefins conversion reaction include “small” and“medium” pore molecular sieves. “Small pore” molecular sieves aredefined as molecular sieves with pores having a diameter of less thanabout 5.0 Angstroms. “Medium pore” molecular sieves are defined asmolecular sieves with pores having a diameter from about 5.0 to about10.0 Angstroms.

[0065] Useful zeolitic molecular sieves include, but are not limited to,mordenite, chabazite, erionite, ZSM-5, ZSM-34, ZSM-48 and mixturesthereof. Methods of making these molecular sieves are known in the artand need not be discussed here. Structural types of small pore molecularsieves that are suitable for use in this invention include AEI, AFT,APC, ATN, ATT, ATV, AWW, BIK, CAS, CHA, CHI, DAC, DDR, EDI, ERI, GOO,KFI, LEV, LOV, LTA, MON, PAU, PHI, RHO, ROG, THO, and substituted formsthereof. Structural types of medium pore molecular sieves that aresuitable for use in this invention include MFI, MEL, MTW, EUO, MTT, HEU,FER, AFO, AEL, TON, and substituted forms thereof.

[0066] Silicoaluminophosphates (“SAPOs”) are one group of non-zeoliticmolecular sieves that are useful in an oxygenate to olefins conversionreaction. SAPOs comprise a three-dimensional microporous crystalframework structure of [SiO₂], [AlO₂ and [PO₂] tetrahedral units. Theway Si is incorporated into the structure is determined by ²⁹Si MAS NMR.See Blackwell and Patton, J. Phys. Chem., 92, 3965 (1988). The desiredSAPO molecular sieves will exhibit one or more peaks in the ²⁹Si MASNMR, with a chemical shift (Si)] in the range of −88 to −96 ppm and witha combined peak area in that range of at least 20% of the total peakarea of all peaks with a chemical shift (Si)] in the range of −88 ppm to−115 ppm, where the (Si)] chemical shifts refer to externaltetramethylsilane (TMS).

[0067] It is desired that the silicoaluminophosphate molecular sieveused in such a process have a relatively low Si/Al₂ ratio. In general,the lower the Si/Al₂ ratio, the lower the C₁-C₄ saturates selectivity,particularly propane selectivity. A Si/Al₂ ratio of less than 0.65 isdesirable, with a Si/Al₂ ratio of not greater than 0.40 being preferred,and a SiAl₂ ratio of not greater than 0.32 being particularly preferred.

[0068] Silicoaluminophosphate molecular sieves are generally classifiedas being microporous materials having 8, 10, or 12 membered ringstructures. These ring structures can have an average pore size rangingfrom about 3.5-15 angstroms. Preferred are the small pore SAPO molecularsieves having an average pore size ranging from about 3.5 to 5angstroms, more preferably from 4.0 to 5.0 angstroms. These pore sizesare typical of molecular sieves having 8 membered rings.

[0069] In general, silicoaluminophosphate molecular sieves comprise amolecular framework of corner-sharing [SiO₂], [AlO₂), and [PO₂]tetrahedral units. This type of framework is effective in convertingvarious oxygenates into olefin products.

[0070] Suitable silicoaluminophosphate molecular sieves for use in anoxygenate to olefin conversion process include SAPO-5, SAPO-8, SAPO-11,SAPO-16, SAPO-17, SAPO-18, SAPO-20, SAPO-31, SAPO-34, SAPO-35, SAPO-36,SAPO-37, SAPO-40, SAPO-41, SAPO-42, SAPO-44, SAPO-47, SAPO-56, the metalcontaining forms thereof, and mixtures thereof. Preferred are SAPO-18,SAPO-34, SAPO-35, SAPO-44, and SAPO-47, particularly SAPO-18 andSAPO-34, including the metal containing forms thereof, and mixturesthereof. As used herein, the term mixture is synonymous with combinationand is considered a composition of matter having two or more componentsin varying proportions, regardless of their physical state.

[0071] Additional olefin-forming molecular sieve materials may be mixedwith the silicoaluminophosphate catalyst if desired. Several types ofmolecular sieves exist, each of which exhibit different properties.Structural types of small pore molecular sieves that are suitable foruse in this invention include AEI, AFT, APC, ATN, ATT, ATV, AWW, BIK,CAS, CHA, CHI, DAC, DDR, EDI, ERI, GOO, KFI, LEV, LOV, LTA, MON, PAU,PHI, RHO, ROG, THO, and substituted forms thereof. Structural types ofmedium pore molecular sieves that are suitable for use in this inventioninclude MFI, MEL, MTW, EUO, MTT, HEU, FER, AFO, AEL, TON, andsubstituted forms thereof. Preferred molecular sieves that may becombined with a silicoaluminophosphate catalyst include ZSM-5, ZSM-34,erionite, and chabazite.

[0072] Substituted SAPOs form a class of molecular sieves known as“MeAPSOs,” which are also useful in the present invention. Processes formaking MeAPSOs are known in the art. SAPOs with substituents, such asMeAPSOs, also may be suitable for use in the present invention. Suitablesubstituents, “Me,” include, but are not necessarily limited to, nickel,cobalt, manganese, zinc, titanium, strontium, magnesium, barium, andcalcium. The substituents may be incorporated during synthesis of theMeAPSOs. Alternately, the substituents may be incorporated aftersynthesis of SAPOs or MeAPSOs using many methods. These methods include,but are not necessarily limited to, ion-exchange, incipient wetness, drymixing, wet mixing, mechanical mixing, and combinations thereof.

[0073] Desired MeAPSOs are small pore MeAPSOs having pore size smallerthan about 5 Angstroms. Small pore MeAPSOs include, but are notnecessarily limited to, NiSAPO-34, CoSAPO-34, NiSAPO-17, CoSAPO-17, andmixtures thereof.

[0074] Aluminophosphates (ALPOs) with substituents, also known as“MeAPOs,” are another group of molecular sieves that may be suitable foruse in an oxygenate to olefin conversion reaction, with desired MeAPOsbeing small pore MeAPOs. Processes for making MeAPOs are known in theart. Suitable substituents include, but are not necessarily limited to,nickel, cobalt, manganese, zinc, titanium, strontium, magnesium, barium,and calcium. The substituents may be incorporated during synthesis ofthe MeAPOs. Alternately, the substituents may be incorporated aftersynthesis of ALPOs or MeAPOs using many methods. The methods include,but are not necessarily limited to, ion-exchange, incipient wetness, drymixing, wet mixing, mechanical mixing, and combinations thereof.

[0075] The molecular sieve may also be incorporated into a solidcomposition, preferably solid particles, in which the molecular sieve ispresent in an amount effective to catalyze the desired conversionreaction. The solid particles may include a catalytically effectiveamount of the molecular sieve and matrix material, preferably at leastone of a filler material and a binder material, to provide a desiredproperty or properties, e.g., desired catalyst dilution, mechanicalstrength and the like, to the solid composition. Such matrix materialsare often to some extent porous in nature and often have somenonselective catalytic activity to promote the formation of undesiredproducts and may or may not be effective to promote the desired chemicalconversion. Such matrix, e.g., filler and binder, materials include, forexample, synthetic and naturally occurring substances, metal oxides,clays, silicas, aluminas, silica-aluminas, silica-magnesias,silica-zirconias, silica-thorias, silica-beryllias, silica-titanias,silica-alumina-thorias, silica-aluminazirconias, and mixtures of thesematerials.

[0076] The solid catalyst composition preferably comprises about 1% toabout 99%, more preferably about 5% to about 90%, and still morepreferably about 10% to about 80%, by weight of molecular sieve; and anamount of about 1% to about 99%, more preferably about 5% to about 90%,and still more preferably about 10% to about 80%, by weight of matrixmaterial.

[0077] The preparation of solid catalyst compositions, e.g., solidparticles, comprising the molecular sieve and matrix material, isconventional and well known in the art and, therefore, is not discussedin detail here.

[0078] The catalyst may further contain binders, fillers, or othermaterial to provide better catalytic performance, attrition resistance,regenerability, and other desired properties. Desirably, the catalyst isfluidizable under the reaction conditions. The catalyst should haveparticle sizes of from about 20μ to about 3,000μ, desirably from about30μ to about 200μ, and more desirably from about 50μ to about 150μ. Thecatalyst may be subjected to a variety of treatments to achieve thedesired physical and chemical characteristics. Such treatments include,but are not necessarily limited to, calcination, ball milling, milling,grinding, spray drying, hydrothermal treatment, acid treatment, basetreatment, and combinations thereof.

[0079] Desirably, in an oxygenate to olefin conversion reactionconducted in the hydrocarbon conversion apparatus of the presentinvention employs a gas superficial velocity in the riser reactors ofgreater than 1 meter per second (m/s). As used herein and in the claims,the term, “gas superficial velocity,” is defined as the volumetric flowrate of vaporized feedstock, and any diluent, divided by the reactorcross-sectional area. Because the oxygenate is converted to a productincluding a light olefin while flowing through the reactor, the gassuperficial velocity may vary at different locations within the reactordepending on the total number of moles of gas present and the crosssection of a particular location in the reactor, temperature, pressure,and other relevant reaction parameters. The gas superficial velocity,including any diluents present in the feedstock, is maintained at a rategreater than 1 meter per second (m/s) at any point in the reactor.Desirably, the gas superficial velocity is greater than about 2 m/s.More desirably, the gas superficial velocity is greater than about 2.5m/s. Even more desirably, the gas superficial velocity is greater thanabout 4 m/s. Most desirably, the gas superficial velocity is greaterthan about 8 m/s.

[0080] Maintaining the gas superficial velocity at these rates increasesthe approach to plug flow behavior of the gases flowing in the riserreactors. As the gas superficial velocity increases above 1 m/s, areduction in axial diffusion or back mixing of the gases results from areduction in internal recirculation of solids, which carry gas withthem. (Ideal plug flow behavior occurs when elements of the homogeneousfluid reactant move through a reactor as plugs moving parallel to thereactor axis). Minimizing the back mixing of the gases in the reactorincreases the selectivity to the desired light olefins in the oxygenateconversion reaction.

[0081] When the gas superficial velocity approaches 1 m/s or higher, asubstantial portion of the catalyst in the reactor may be entrained withthe gas exiting the riser reactors. At least a portion of the catalystexiting the riser reactors is recirculated to recontact the feed throughthe catalyst return.

[0082] Desirably, the rate of catalyst, comprising molecular sieve andany other materials such as binders, fillers, etc., recirculated torecontact the feed is from about 1 to about 100 times, more desirablyfrom about 10 to about 80 times, and most desirably from about 10 toabout 50 times the total feed rate, by weight, of oxygenates to thereactor.

[0083] The temperature useful to convert oxygenates to light olefinsvaries over a wide range depending, at least in part, on the catalyst,the fraction of regenerated catalyst in a catalyst mixture, and theconfiguration of the reactor apparatus and the reactor. Although theseprocesses are not limited to a particular temperature, best results areobtained if the process is conducted at a temperature from about 200° C.to about 700° C., desirably from about 250° C. to about 600° C., andmost desirably from about 300° C. to about 500° C. Lower temperaturesgenerally result in lower rates of reaction, and the formation rate ofthe desired light olefin products may become markedly slower. However,at temperatures greater than 700° C., the process may not form anoptimum amount of light olefin products, and the rate at which coke andlight saturates form on the catalyst may become too high.

[0084] Light olefins will form—although not necessarily in optimumamounts—at a wide range of pressures including, but not limited to,pressures from about 0.1 kPa to about 5 MPa. A desired pressure is fromabout 5 kPa to about 1 MPa and most desirably from about 20 kPa to about500 kPa. The foregoing pressures do not include that of a diluent, ifany, and refer to the partial pressure of the feed as it relates tooxygenate compounds and/or mixtures thereof. Pressures outside of thestated ranges may be used and are not excluded from the scope of theinvention. Lower and upper extremes of pressure may adversely affectselectivity, conversion, coking rate, and/or reaction rate; however,light olefins will still form and, for that reason, these extremes ofpressure are considered part of the present invention.

[0085] A wide range of WHSV's for the oxygenate conversion reaction,defined as weight of total oxygenate fed to the riser reactors per hourper weight of molecular sieve in the catalyst in the riser reactors,function with the present invention. The total oxygenate fed to theriser reactors includes all oxygenate in both the vapor and liquidphase. Although the catalyst may contain other materials, which act asinerts, fillers or binders, the WHSV is calculated using only the weightof molecular sieve in the catalyst in the riser reactors. The WHSV isdesirably high enough to maintain the catalyst in a fluidized stateunder the reaction conditions and within the reactor configuration anddesign. Generally, the WHSV is from about 1 hr⁻¹ to about 5000 hr⁻¹,desirably from about 2 hr⁻¹ to about 3000 hr⁻¹, and most desirably fromabout 5 hr⁻¹ to about 1500 hr⁻¹. The applicants have discovered thatoperation of the oxygenate to olefin conversion reaction at a WHSVgreater than 20 hr⁻¹ reduces the methane content in the product slate ofthe conversion reaction. Thus, the conversion reaction is desirablyoperated at a WHSV of at least about 20 hr⁻¹. For a feed comprisingmethanol, dimethyl ether, or mixtures thereof, the WHSV is desirably atleast about 20 hr⁻¹ and more desirably from about 20 hr⁻¹ to about 300hr⁻¹.

[0086] It is particularly preferred that the reaction conditions formaking olefins from an oxygenate comprise a WHSV of at least about 20hr⁻¹ and a Temperature Corrected Normalized Methane Selectivity (TCNMS)of less than about 0.016. As used herein, TCNMS is defined as theNormalized Methane Selectivity (NMS) when the temperature is less than400° C. The NMS is defined as the methane product yield divided by theethylene product yield wherein each yield is measured on or is convertedto a weight % basis. When the temperature is 400° C. or greater, theTCNMS is defined by the following equation, in which T is the averagetemperature within the reactor in ° C.:${TCNMS} = {\frac{NMS}{1 + \left( {\left( {\left( {T - 400} \right)/400} \right) \times 14.84} \right)}.}$

[0087] Oxygenate conversion should be maintained sufficiently high toavoid the need for commercially unacceptable levels of feed recycling.While 100% oxygenate conversion is desired for the purpose of completelyavoiding feed recycle, a reduction in unwanted by-products is observedfrequently when the conversion is about 98% or less. Since recycling upto as much as about 50% of the feed is commercially acceptable,conversion rates from about 50% to about 98% are desired. Conversionrates may be maintained in this range—50% to about 98%—using a number ofmethods familiar to persons of ordinary skill in the art. Examplesinclude, but are not necessarily limited to, adjusting one or more ofthe following: reaction temperature; pressure; flow rate (weight hourlyspace velocity and/or gas superficial velocity); catalyst recirculationrate; reactor apparatus configuration; reactor configuration; feedcomposition; amount of liquid feed relative to vapor feed (as will bediscussed below); amount of recirculated catalyst; degree of catalystregeneration; and other parameters, which affect the conversion.

[0088] During the conversion of oxygenates to light olefins,carbonaceous deposits accumulate on the catalyst used to promote theconversion reaction. At some point, the build up of these carbonaceousdeposits causes a reduction in the capability of the catalyst to convertthe oxygenate feed to light olefins. At this point, the catalyst ispartially deactivated. When a catalyst can no longer convert anoxygenate to an olefin product, the catalyst is considered to be fullydeactivated. As an optional step in an oxygenate to olefin conversionreaction, a portion of the catalyst is withdrawn from the reactor and atleast a portion of the portion removed from the reactor is partially, ifnot fully, regenerated in a regeneration apparatus, such as regenerationapparatus 80 as shown in FIG. 4. By regeneration, it is meant that thecarbonaceous deposits are at least partially removed from the catalyst.Desirably, the portion of the catalyst withdrawn from the reactor is atleast partially deactivated. The remaining portion of the catalyst inthe reactor is re-circulated without regeneration, as described above.The regenerated catalyst, with or without cooling, is then returned tothe reactor. Desirably, the rate of withdrawing the portion of thecatalyst for regeneration is from about 0.1% to about 99% of the rate ofthe catalyst exiting the reactor. More desirably, the rate is from about0.2% to about 50%, and, most desirably, from about 0.5% to about 5%.

[0089] Desirably, a portion of the catalyst, comprising molecular sieveand any other materials such as binders, fillers, etc., is removed fromthe reactor for regeneration and recirculation back to the reactor at arate of from about 0.1 times to about 10 times, more desirably fromabout 0.2 to about 5 times, and most desirably from about 0.3 to about 3times the total feed rate of oxygenates to the reactor. These ratespertain to the catalyst containing molecular sieve only, and do notinclude non-reactive solids. The rate of total solids, i.e., catalystand non-reactive solids, removed from the reactor for regeneration andrecirculation back to the reactor will vary these rates in directproportion to the content of non-reactive solids in the total solids.

[0090] Desirably, the catalyst regeneration is carried out in aregeneration apparatus in the presence of a gas comprising oxygen orother oxidants. Examples of other oxidants include, but are notnecessarily limited to, singlet O₂, O₃, SO₃, N₂O, NO, NO₂, N₂O₅, andmixtures thereof. Air and air diluted with nitrogen or CO₂ are desiredregeneration gases. The oxygen concentration in air may be reduced to acontrolled level to minimize overheating of, or creating hot spots in,the regenerator. The catalyst may also be regenerated reductively withhydrogen, mixtures of hydrogen and carbon monoxide, or other suitablereducing gases.

[0091] The catalyst may be regenerated in any number of methods—batch,continuous, semi-continuous, or a combination thereof. Continuouscatalyst regeneration is a desired method. Desirably, the catalyst isregenerated to a level of remaining coke from about 0.01 wt % to about15 wt % of the weight of the catalyst.

[0092] The catalyst regeneration temperature should be from about 250°C. to about 750° C., and desirably from about 500° C. to about 700° C.Because the regeneration reaction takes place at a temperatureconsiderably higher than the oxygenate conversion reaction, it may bedesirable to cool at least a portion of the regenerated catalyst to alower temperature before it is sent back to the reactor. A heatexchanger, not shown, located external to the regeneration apparatus maybe used to remove some heat from the catalyst after it has beenwithdrawn from the regeneration apparatus. When the regenerated catalystis cooled, it is desirable to cool it to a temperature, which is fromabout 200° C. higher to about 200° C. lower than the temperature of thecatalyst withdrawn from the reactor. More desirably, the regeneratedcatalyst is cooled to a temperature from about 10° C. to about 200° C.lower than the temperature of the catalyst withdrawn from the reactor.This cooled catalyst then may be returned to either some portion of thereactor, the regeneration apparatus, or both. When the regeneratedcatalyst from the regeneration apparatus is returned to the reactor, itmay be returned to any portion of the reactor. It may be returned to thecatalyst containment area to await contact with the feed, the separationzone to contact products of the feed or a combination of both.

[0093] Desirably, catalyst regeneration is carried out at leastpartially deactivated catalyst that has been stripped of most of readilyremovable organic materials (organics) in a stripper or strippingchamber first. This stripping is achieved by passing a stripping gasover the spent catalyst at an elevated temperature. Gases suitable forstripping include steam, nitrogen, helium, argon, methane, CO₂, CO,hydrogen, and mixtures thereof. A preferred gas is steam. Gas hourlyspace velocity (GHSV, based on volume of gas to volume of catalyst andcoke) of the stripping gas is from about 0.1 h⁻¹ to about 20,000 h⁻¹.Acceptable temperatures of stripping are from about 250° C. to about750° C., and desirably from about 350° C. to about 675° C.

[0094] The process of making the preferred olefin product in thisinvention can include the additional step of making the oxygenatecompositions from hydrocarbons such as oil, coal, tar sand, shale,biomass and natural gas. Processes for making the compositions are knownin the art. These processes include fermentation to alcohol or ether,making synthesis gas, then converting the synthesis gas to alcohol orether. Synthesis gas is produced by known processes such as steamreforming, autothermal reforming and partial oxidization.

[0095] One skilled in the art will also appreciate that the olefinsproduced by the oxygenate-to-olefin conversion reaction of the presentinvention optionally are polymerized to form polyolefins, particularlypolyethylene and polypropylene. Processes for forming polyolefins fromolefins are known in the art. Catalytic processes are preferred.Particularly preferred are metallocene, Ziegler/Natta and acid catalyticsystems. See, for example, U.S. Pat. Nos. 3,258,455; 3,305,538;3,364,190; 5,892,079; 4,659,685; 4,076,698; 3,645,992; 4,302,565; and4,243,691, the catalyst and process descriptions of each being expresslyincorporated herein by reference. In general, these processes involvecontacting the olefin product with a polyolefin-forming catalyst at apressure and temperature effective to form the polyolefin product.

[0096] A preferred polyolefin-forming catalyst is a metallocenecatalyst. The preferred temperature range of operation is between 50° C.and 240° C. and the reaction is carried out at low, medium or highpressure, being anywhere from 1 bar to 200 bars. For processes carriedout in solution, an inert diluent optionally is used, and the preferredoperating pressure range is between 10 and 150 bars, with a preferredtemperature between 120° C. and 230° C. For gas phase processes, it ispreferred that the temperature generally be from 60° C. to 160° C., andthat the operating pressure be from 5 bars to 50 bars.

[0097] In addition to polyolefins, numerous other olefin derivatives maybe formed from the olefins produced by the process of the presentinvention or olefins recovered therefrom. These include, but are notlimited to, aldehydes, alcohols, acetic acid, linear alpha olefins,vinyl acetate, ethylene dichloride and vinyl chloride, ethylbenzene,ethylene oxide, ethylene glycol, cumene, isopropyl alcohol, acrolein,allyl chloride, propylene oxide, acrylic acid, ethylene-propylenerubbers, and acrylonitrile, and trimers and dimers of ethylene,propylene or butylenes. The processes of manufacturing these derivativesare well known in the art, and therefore are not discussed here.

[0098] Persons of ordinary skill in the art will recognize that manymodifications may be made to the present invention without departingfrom the spirit and scope of the present invention. The embodimentsdescribed herein are meant to be illustrative only and should not betaken as limiting the invention, which is defined by the followingclaims.

1. A hydrocarbon conversion apparatus, comprising: (a) a plurality ofriser reactors, each having a first end into which a catalyst can befed, and a second end through which the catalyst can exit; (b) aseparation zone having a plurality of inlets, the separation zone beingprovided to separate the catalyst from products of a reaction conductedin the hydrocarbon conversion apparatus; (c) a plurality of deviatingmembers, each deviating member being in fluid communication between thesecond end of a respective riser reactor and a respective inlet of theseparation zone; (d) a catalyst retention zone provided to containcatalyst, which can be fed to the riser reactors; and (e) a catalystreturn in fluid communication between the separation zone and thecatalyst retention zone.
 2. The apparatus of claim 1, wherein each ofthe plurality of riser reactors includes a center axis extending betweenthe first and second ends thereof, and wherein the plurality of inletsare not oriented along the center axes of the riser reactors.
 3. Thehydrocarbon conversion apparatus of claim 2, wherein the apparatusfurther comprises: (f) a plurality of arms, each arm in fluidcommunication between the catalyst retention zone and the first end of arespective riser reactor.
 4. The hydrocarbon conversion apparatus ofclaim 2, wherein the apparatus further comprises: (f) a feed distributorincluding a plurality of feed heads positioned adjacent to the firstends of the riser reactors.
 5. The hydrocarbon conversion apparatus ofclaim 2, wherein the hydrocarbon conversion apparatus includes two riserreactors.
 6. The hydrocarbon conversion apparatus of claim 2, whereinthe hydrocarbon conversion apparatus includes three riser reactors. 7.The hydrocarbon conversion apparatus of claim 2, wherein the hydrocarbonconversion apparatus includes four riser reactors.
 8. The hydrocarbonconversion apparatus of claim 2, wherein the hydrocarbon conversionapparatus includes five riser reactors.
 9. The hydrocarbon conversionapparatus of claim 2, wherein the hydrocarbon conversion apparatusincludes six riser reactors.
 10. The hydrocarbon conversion apparatus ofclaim 2, wherein the hydrocarbon conversion apparatus includes more thansix riser reactors.
 11. The hydrocarbon conversion apparatus of claim 4,wherein the feed distributor provides feed to each of the riser reactorsin substantially equal streams through the feed heads.
 12. Thehydrocarbon conversion apparatus of claim 2, wherein the apparatusfurther comprises: (f) a fluidizing agent distributor in fluidcommunication with the catalyst retention zone, the fluidizing agentdistributor being provided to feed a fluidizing agent to the catalystretention zone to fluidize catalyst contained in the catalyst retentionzone.
 13. The hydrocarbon conversion apparatus of claim 12, wherein theapparatus further comprises: (g) a disperser positioned in the catalystretention zone, the disperser being provided to disperse the fluidizingagent in the catalyst retention zone to fluidize the catalyst.
 14. Thehydrocarbon conversion apparatus of claim 13, wherein the disperser is adevice selected from the group consisting of a grid, a screen and aperforated plate.
 15. The hydrocarbon conversion apparatus of claim 2,wherein the catalyst return is positioned centrally to the riserreactors.
 16. The hydrocarbon conversion apparatus of claim 2, whereinthe hydrocarbon conversion apparatus includes a plurality of catalystreturns.
 17. The hydrocarbon conversion apparatus of claim 16, whereinthe apparatus further comprises: (f) a flow control device positioned onat least one of the catalyst returns.
 18. The hydrocarbon conversionapparatus of claim 16, wherein the apparatus further comprises: (f) aflow control device positioned on each of the plurality of catalystreturns.
 19. The hydrocarbon conversion apparatus of claim 2, whereinthe separation zone further comprises a quiescent zone in which catalystcan be retained until the catalyst moves from the separation zone to thecatalyst return.
 20. The hydrocarbon conversion apparatus of claim 2,wherein the apparatus further comprises: (f) a catalyst regenerator influid communication with the hydrocarbon conversion apparatus.
 21. Thehydrocarbon conversion apparatus of claim 20, wherein the apparatusfurther comprises: (g) a catalyst stripper in fluid communication withthe hydrocarbon conversion apparatus and the catalyst regenerator. 22.The hydrocarbon conversion apparatus of claim 2, wherein the apparatusfurther comprises: (f) at least one separation device positioned in theseparation zone.
 23. The hydrocarbon conversion apparatus of claim 22,wherein the at least one separation device is selected from groupconsisting of a cyclonic separator, a filter, an impingement device andcombinations thereof.
 24. The hydrocarbon conversion apparatus of claim2, wherein each of the riser reactors has a cross sectional area of nogreater than 12 m².
 25. The hydrocarbon conversion apparatus of claim 2,wherein each of the riser reactors has a cross sectional area of nogreater than 7 m².
 26. The hydrocarbon conversion apparatus of claim 2,wherein each of the riser reactors has a cross sectional area or nogreater than 3.5 m².
 27. The hydrocarbon conversion apparatus of claim2, wherein each of the riser reactors has a height of from 10 meters to70 meters.
 28. The hydrocarbon conversion apparatus of claim 2, whereineach of the riser reactors has a width of from 1 meter to 3 meters. 29.The hydrocarbon conversion apparatus of claim 2, wherein each of theriser reactors has a cross sectional area and the cross sectional areaof one of the riser reactors varies by no more than 20% from the crosssectional area of another of the riser reactors.
 30. The hydrocarbonconversion apparatus of claim 29, wherein the cross sectional area ofone of the riser reactors varies by no more than 10% from the crosssectional area of another of the riser reactors.
 31. The hydrocarbonconversion apparatus of claim 30, wherein the cross sectional area ofone of the riser reactors varies by no more than 1% from the crosssectional area of another of the riser reactors.
 32. The hydrocarbonconversion apparatus of claim 2, wherein at least one of the deviatingmembers comprises a tubular member providing a deviation angle of atleast 45 degrees.
 33. The hydrocarbon conversion apparatus of claim 2,wherein at least one of the deviating members comprises a tubular memberproviding a deviation angle of at least 90 degrees.
 34. The hydrocarbonconversion apparatus of claim 2, wherein at least one of the deviatingmembers comprises a 90 degree elbow.
 35. The hydrocarbon conversionapparatus of claim 2, wherein the apparatus further comprises: (f) asecond plurality of riser reactors, each of the second plurality ofriser reactors having first and second ends and a second center axis,wherein the separation zone includes a plurality of second inlets, eachsecond inlet being oriented along a respective second center axis.
 36. Ahydrocarbon conversion apparatus, comprising: (a) a plurality of riserreactors, each having a first end into which a catalyst can be fed and asecond end through which the catalyst can exit the riser reactor; (b) aseparation zone provided to separate the catalyst from products of areaction conducted in the hydrocarbon conversion apparatus, theseparation zone including a plurality of inlets; (c) a plurality ofdeviating members, each deviating member oriented to deviate a flow ofmaterial from the second end of a respective riser reactor toward arespective inlet; and (d) at least one catalyst return in fluidcommunication with the separation zone and the first ends of the riserreactors, the catalyst return being provided to transfer the catalystfrom the separation zone to the first ends of the riser reactors. 37.The hydrocarbon conversion apparatus of claim 36, wherein each of theriser reactors has a height of from 10 meters to 70 meters.
 38. Thehydrocarbon conversion apparatus of claim 36, wherein each of the riserreactors has a width of from one meter to three meters.
 39. Thehydrocarbon conversion apparatus of claim 36, wherein each of the riserreactors has a cross sectional area of no greater than 12 m².
 40. Thehydrocarbon conversion apparatus of claim 36, wherein each of the riserreactors has a cross sectional area of no greater than 7 m².
 41. Thehydrocarbon conversion apparatus of claim 36, wherein each of the riserreactors has a cross sectional area or no greater than 3.5 m².
 42. Thehydrocarbon conversion apparatus of claim 36, wherein each of the riserreactors has a cross sectional area and the cross sectional area of oneof the riser reactors varies by no more than 20% from the crosssectional area of another of the riser reactors.
 43. The hydrocarbonconversion apparatus of claim 42, wherein the cross sectional area ofone of the riser reactors varies by no more than 10% from the crosssectional area of another of the riser reactors.
 44. The hydrocarbonconversion apparatus of claim 43, wherein the cross sectional area ofone of the riser reactors varies by no more than 1% from the crosssectional area of another of the riser reactors.
 45. The hydrocarbonconversion apparatus of claim 36, wherein the apparatus furthercomprises: (e) a catalyst retention zone provided to contain catalystwhich can be fed to the riser reactors.
 46. The hydrocarbon conversionapparatus of claim 45, wherein the apparatus further comprises: (f) afeed distributor including a plurality of feed heads, each head beingpositioned adjacent to the first end of a respective riser reactor. 47.The hydrocarbon conversion apparatus of claim 36, wherein the apparatusfurther comprises: (e) a fluidizing agent distributor in fluidcommunication with the catalyst return, the fluidizing agent distributorbeing provided to feed a fluidizing agent to the catalyst return tofluidize catalyst contained in the catalyst return.
 48. The hydrocarbonconversion apparatus of claim 47, wherein the apparatus furthercomprises: (f) a regeneration apparatus in fluid communication with thehydrocarbon conversion apparatus.
 49. The hydrocarbon conversionapparatus of claim 48, wherein the apparatus further comprises: (g) acatalyst stripper in fluid communication with the regenerationapparatus.
 50. The hydrocarbon conversion apparatus of claim 36, whereinat least one of the deviating members comprises a tubular memberproviding a deviation angle of at least 45 degrees.
 51. The hydrocarbonconversion apparatus of claim 36, wherein at least one of the deviatingmembers comprises a tubular member providing a deviation angle of atleast 90 degrees.
 52. The hydrocarbon conversion apparatus of claim 36,wherein at least one of the deviating members comprises a 90 degreeelbow.
 53. The hydrocarbon conversion apparatus of claim 36, wherein theapparatus further comprises: (e) a second plurality of riser reactors,each of the second plurality of riser reactors having a first end intowhich a catalyst can be fed, and a second end through which the catalystcan exit the riser reactor and directly enter the separation zone.
 54. Ahydrocarbon conversion process, comprising the steps of: (a) contactinga fluidizable catalyst with a fluidizing agent to fluidize thefluidizable catalyst; (b) feeding the catalyst and a feed to a pluralityof riser reactors, the plurality of riser reactors being part of asingle hydrocarbon conversion apparatus; (c) contacting the feed withthe catalyst in the plurality of riser reactors under conditionseffective to convert the feed to a product; (d) directing the productand the catalyst through a plurality of deviating members each deviatingmember being positioned to deviate a flow of the product and thecatalyst from an outlet of a respective riser reactor to a separationzone; (e) separating the catalyst from the product in the separationzone, the separation zone being in fluid communication with theplurality of deviating members; (f) returning the catalyst from theseparation zone to the plurality of riser reactors; and (g) repeatingsteps (a) to (f).
 55. The process of claim 54, wherein the feed is fedto each of the plurality of riser reactors in a substantially equalamount.
 56. The process of claim 54, wherein the feed is fed to each ofthe plurality of riser reactors such that the flow of feed to eachreactor varies by no more than 25%, by volume rate, from one riserreactor to another riser reactor.
 57. The process of claim 56, whereinthe feed is fed to each of the plurality of riser reactors such that theflow of feed to each reactor varies by no more than 10%, by volume rate,from one riser reactor to another riser reactor.
 58. The process ofclaim 57, wherein the feed is fed to each of the plurality of riserreactors such that the flow of feed to each reactor varies by no morethan 1%, by volume rate, from one riser reactor to another riserreactor.
 59. The process of claim 54, wherein the feed is fed to each ofthe plurality of riser reactors such that the flow of feed to eachreactor varies by no more than 25%, by mass percent for each componentin the feed, from one riser reactor to another riser reactor.
 60. Theprocess of claim 59, wherein the feed is fed to each of the plurality ofriser reactors such that the flow of feed to each reactor varies by nomore than 10%, by mass percent for each component in the feed, from oneriser reactor to another riser reactor.
 61. The process of claim 60,wherein the feed is fed to each of the plurality of riser reactors suchthat the flow of feed to each reactor varies by no more than 1%, by masspercent for each component in the feed, from one riser reactor toanother riser reactor.
 62. The process of claim 54, wherein thefluidizing agent is selected from the group consisting of nitrogen,steam, carbon dioxide, hydrocarbons and air.
 63. The process of claim54, wherein the catalyst is separated from the product with a separationdevice selected from the group consisting of cyclonic separators,filters, screens, impingement devices, plates, cones and combinationsthereof.
 64. The process of claim 54, wherein the returning furthercomprises directing the catalyst from the separation zone to a catalystreturn, which is in fluid communication with the separation zone and aplurality of arms.
 65. The process of claim 64, wherein the returningfurther comprises directing the catalyst through the plurality of armsand to an inlet on each respective riser reactors.
 66. The process ofclaim 64, wherein the catalyst is contacted with the fluidizing agent tofluidize the fluidizable catalyst in the catalyst return, in a catalystretention zone or a combination of the catalyst return and the catalystretention zone.
 67. The process of claim 66, wherein the process furthercomprises the steps of: (h) regenerating at least a portion of thecatalyst in a catalyst regenerator after separating the catalyst fromthe products to produce a regenerated catalyst; and (i) returning theregenerated catalyst to at least one of the separation zone, thecatalyst return, and the catalyst retention zone.
 68. The process ofclaim 67, wherein the process further comprises the step of: (j)stripping the at least a portion of the catalyst prior to regeneratingthe at least a portion of the catalyst.
 69. The process of claim 54,wherein each of the riser reactors has a height of from 10 meters to 70meters.
 70. The process of claim 54, wherein each of the riser reactorshas a width of from 1 meter to 3 meters.
 71. The process of claim 54,wherein each of the riser reactors has a cross sectional area of nogreater than 12 m².
 72. The process of claim 54, wherein each of theriser reactors has a cross sectional area of no greater than 7 m². 73.The process of claim 54, wherein each of the riser reactors has a crosssectional area or no greater than 3.5 m².
 74. The process of claim 54,wherein each of the riser reactors has a cross sectional area and thecross sectional area of one of the riser reactors varies by no more than20% from the cross sectional area of another of the riser reactors. 75.The process of claim 74, wherein the cross sectional area of one of theriser reactors varies by no more than 10% from the cross sectional areaof another of the riser reactors.
 76. The process of claim 75, whereinthe cross sectional area of one of the riser reactors varies by no morethan 1% from the cross sectional area of another of the riser reactors.77. The process of claim 54, wherein the hydrocarbon conversion processis a reaction selected from the group consisting of an olefininterconversion reaction, an oxygenate to olefin conversion reaction, anoxygenate to gasoline conversion reaction, malaeic anhydrideformulation, vapor phase methanol synthesis, phthalic anhydrideformulation, a Fischer Tropsch reaction, and acrylonitrile formulation.78. The process of claim 54, wherein the hydrocarbon conversion processis an oxygenate to olefin conversion reaction.
 79. The process of claim78, wherein the catalyst is a silicoaluminophosphate catalyst.
 80. Theprocess of claim 79, wherein the feed is selected from the group ofmethanol; ethanol; n-propanol; isopropanol; C₄-C₁₀ alcohols; methylethyl ether; dimethyl ether; diethyl ether; di-isopropyl ether; methylformate; formaldehyde; di-methyl carbonate; methyl ethyl carbonate;acetone; and mixtures thereof.
 81. The process of claim 54, wherein atleast one of the deviating members deviates the flow of the product andthe catalyst by at least 45 degrees.
 82. The process of claim 54,wherein at least one of the deviating members deviates the flow of theproduct and the catalyst by at least 90 degrees.
 83. The process ofclaim 54, wherein at least one of the deviating members comprises a 90degree elbow.
 84. A hydrocarbon conversion apparatus, comprising: (a) aplurality of riser reactors, each having a first end for receivingcatalyst and a second end through which the catalyst and a product canexit; (b) a plurality of deviating members, each associated with arespective riser reactor; (c) a separation zone having at least one sideand provided to separate the catalyst from the product, wherein theseparation zone includes a plurality of inlets, each inlet beingassociated with a respective deviating member, wherein the inlets areoriented on the side of the separation zone; and (d) a catalyst returncoupled to a plurality of arms, the catalyst return and arms being influid communication between the separation zone and the first ends ofthe plurality of riser reactors.